Continuous fluid hydroforming



R. E. MacPHERsoN, JR., x-:TAL .2,656,304

`CONTINUOUS FLUID HYDROFORMING Oct. 20, 1953 Filed Feb. 28 1951 dmaondu@ paru. L .@.maqphmmgfn Patented Oct. 20, 1953 2,656,302! y ooNfrINtIo'Us HYDROFORMING assert Matheson, Ji., Crim-fpm, aiiefttb 'Schricker,-zlf1., Roselle, N. 'Jj., ^ass"gno"s 'tb Stand@ are on oeveiopiiient eenparig, s csr'poiatiuii of Delaware Application February z8, '1931, 'se'iiogac (ci. rs'ceiso) 1 i Y i lThe novelty of the present invention fully Edisclosed in the ollowinig `S15ecil'c'ai'ii'tn and tel-aims, considered in `conne'ftion with the acf'cornpanying drawing.

Heretofore and prior to the present invention,

'it was a matter of record and commercial practice to hydroform hydrocarbons, particularly those containing appreciable amounts ofvnaphthenic hydrocarbons in the `presence of solid catalysts and added hydrogen tov form a product of increased aromaticity. It is generally considered that the `main reaction involved is one of dehydrogenation of the naphtloenes to the corresponding aromatics; and `'at the same time the hydrogenation of olenic hydrocarbons that may be present in order to convert them to the corresponding paraiinic hydrocaibons. Furthermore, it "is generally considered that a certain amount of isomerization 'ci olefiiiic and paranic hydrocarbons occurs during this hydroforming operation'. It is also 'gen- 'e'rally considered that isorniza'tio of naphf the-nic com-pounds takes place during the Vhydroforming reaction, as Where ethyl-cyclobentane is converted to methyl-cyclohean'e.

This hydro-forming operation, acccirdiiiffgy to ier-lor practice, was carriedout at elevated ternpe'ratures and pressures; The operation isu-Its `irl the deposition of carbonaeous vI'r'iateris'tl`s' on the solid catalyst, and period-ic regeneration o'f the catalyst `to restore its activity is necessary. Furthermore, the commercial 'opration was carried out in the presence of i''d beds of catalyst disposed in a plurality of reactors operating lh series. Due to the high en'dothermic nature lof the reaction, a severe temperature drop was encountered particularly in the lead reactor; This is a very unsatisfactory condition because the feed to the lead reactor generally had to be heated to a temperature far above that desired in order that the reactants would be at suinci'entl'y high temperature to cause the reaction to proceed at a reasonable rate downstream from the inlet point. This excessivahcating of the feed to the lead reactor resulted in the forma"- tion of inordinately large quantities 'of normal; ly gaseous material and carbon.

The present invention constitutes an improvement over prior practicel in many respects.. n the rst place, the catalyst is disposed in the form of a dense fluidiz'ed bedA in the reactor. This technique, due to the intimate thorouali niiiiing of all portions of th catalyst, tends to iiiaiiitain the 'catalyst bed at a substantially uniform temperature. I

Another advantage of the present improve'- nfients resides in the fact that the operation of hydro'forming may be carried out continuously Without interruption of the productive phase to regenerate the catalyst,- us-ing the saine catalyst,

vtion is to subject hydrocarbons tQ hydro fprviiisiy requiring regeneration with air tr "some other 'a's in prior practice.

Another advantage oi the' present invention resides' in the possibility of ."ibtairiing very hih ctfane rating motor fiifel or aviation gasoline.

The iiiaiii iibject or the present 'iiiveiitiii is jto carry out hydrjfi'i'ing operation under condi- 'tions which afford 'greater ii'eiiibiiity ci opration, more efficient operation, and at the "saine time, important savings in the 'c'ost of peation are secured.

A more speciiic object ci the present under conditions suchA as to produce H y, hating an cctaiiejmimbr 'iip to about 95 erna. other and' ruitirer objects are ci, .prsent invention wiu appear iiciii the `iouoivifi'g more 'detailed description.

N'In the accoinpanyin drawing there 'shown uiasiamiiiaticailyin Fi .(1, a suitable apparatus 4for car`1'^`y`i'ng the `present invention into e ect.

1 Referring iii detail to Fig. '1, feed stoer fits the present systerfthr'ough line I, is piinpedvby piiiiiib 2., through iin s, into jsuitabie iis-ating means 4, which may be an ordinary 'ed cifl. Simultaneously hydrogen from some ,source eil'fs the prsent syst'n [1J-l1v ugh ln` v5, is there? after forced tiiroirgii @interresser e, then passe via iiiiiiv iiiiitb a. flied con a Where it is also re"- jiiated. The heated oii is withdrawn 'from th furnace into line 9 and the heated liydroen is withdrawn weer line m ad m in een aiateiy introduced iiito iiydjriirriiiiiig Vii'eiii-i'tiiji I2 as shown in the drawing.

to a higher 'temperature than the il in ord a prevent thermal cracking of the latter, and in this connection therefore, the heated il and the heated iiyuiiigemcoiitaiiiiiig as are preferably separately introduced into reactor I2',` the fdr-mer Fbelow a gas distributor G and the latter above. The catalyst Qin I 2 is maintained in the form f a dense luidized bed extending from the distributing grid G to an upper dense phase level L. The catalyst itself will be described more fully hereinafter as ,to composition;e For j the present it will simply be stated that the catalyst may be a Lgroip Vvinetal oxide on a suitable carrier and isf rund to pvv'dr df uidizable size, say hating' particle size i fr about 20G-40D iesli or finer. The s'p'erlicial velocity .i am gli of the `oil vapors and the hydrogiil''ntainmg gas in the reactor l? is maintained at a value oflfrom about 3%; to 1 'foot per second. The reactants are maintained in contact with the cat; alyst under conditions o'f temprature, pressure, contact time, etc., more fully described JHereinafter, to effect the desired conversion and ually issue from the dense fiuidie'd bedy to the top of the reactor I2. As usual in this typ cf operation, a iight dispersed phase of catalyst in gasiforrn material exists in the reactor above the level L. The reaction vapors, before issuing from the reactor I2, are forced through one or more gas-solids separating devices I3, for the purpose of separating entrained catalysts, this entrained catalyst being withdrawn through one or more dip pipes d from the separating device and returned to the bed of catalyst C. rThe reactant products are withdrawn from the reactor I2 through line I4 and preferably are passed through conventional oil scrubbing means I5 in which catalyst still contained in the vapors is removed and returned via line I6 to the bed of catalyst C. An alternative method of removing these last traces of catalyst is to eiect a partial condensation of the vapors in S whereby the entrained catalyst is washed out of the said vapors to form an oil and catalyst slurry, which is then returned to the reactor I2 as previously noted. The product substantially free of catalyst, passes from S via line l? in a high pressure separator I 8. A gas rich in hydrogen is recovered overhead Via line I 9 from separator I8 and recycled to hydrogen supply line 5 for further use in the process. A portion of the recycled gas may be continuously or intermittently bled from the system through line 20. Referring again to high pressure separator I8, the crude liquid product is withdrawn from I8 via line 2| and forced through pressure reducing valve 22 and thereafter introduced into fractionating tower 23 where it is subjected to fractionation and distillation to recover the desired product. Light material, that is to say, low boiling material is withdrawn from the fractionator 23 through line 24% and rejected from the system. The hydroformed product is recovered from fractionator 23 through line 25 and collected in receiving drum 26.

It will be understood that a petroleum engineer would understand that in a commercial plant a great deal of accessory apparatus not shown in the drawing would be required for efcient operation. Thus, the drawing does not include such devices as would ordinarily be used, such as ow meters, temperature recording devices, ltering devices, etc.

In order further to illustrate the present invention the following specific examples are set forth.

There is set forth below, an inspection of the naphtha which was used in these runs.

Feed inspection West Texas heavy naphtha, approximately 200-430 F. Boiling range A. P. I. gravity 53.3

Octane No., CFRR 43 Reid vapor pressure 0.7 lbs.

Hydrogen concentration in recycle gas, vol. percent 70-75 Hydrogen partial pressure in reactor,

p. s. i. g 460' Hydrogen consumption SCF/bbl (Z-380) Length of run, hours 210 Deactivation rate, A O. N./hr 0.0

1Prepared by zo-precipitation of (NH4)2 M00; in aqueous solution with A1Cl3.

An inspection of the product revealed the following:

Product A. P. I. gravity 62.2 Yield in vol. per cent C4+ 85.8 Octane No., CFRR 92.0 Reid Vapor pressure 16.00 Aromatics, Vol. per cent 28.00 Olens, vol. per cent 2 Sulfur Trace Naphthenes 2.5% Paraiiins (i5-70% In the below Examples II and III, the same feed was employed, as noted above, as in EX- ample I, but it will be noted that in these examples the catalyst became deactivated since the octane number decreased 0.061 in Run No. -4 (Example II) and 0.065 octane number per hour in Run No. 61-3 (Example III).

1 Prepared by co-precipitaton of (NH4)2 M004 in aqueous solution with AlCla.

EXAMPLE III Run No 61-3 Catalyst 8.5 wt. per cent M003 on alumina.l Temperature, F 930 Pressure, p. s. i. g 500 Feed rate, w/hr./w 0.25 Recycle rate, SCF/bbl 6000-'7000 Hydrogen concentration in recycle gas, vol. per cent 60-65 Hydrogen partial pressure in reactor.

p. s. i. g 270 Hydrogen consumption, SCF/bbl 210-(60) Length of run, hours 210 Deactivation rate, A O. N./hr 0.065

1 Prepared by co-precipitatlon of (NHQZ H004 in aqueous solution with AlClS.

2 Some hydrogen was produced in this run.

It will be noted that in Run 60-3 (Example I), the hydrogen partial pressure was 460 p. s. i. g. and under these conditions the activity of the catalyst was maintained or, in other words, the process of hydroforming was continuous, not requiring interruption to regenerate the catalyst. However, in Run 60-4 (Example II) where the hydrogen partial pressure was 410 p. s. i. g., the octane number of the product decreased at a rate of 0.061 number per hour which, oi course, means that the rcatalyst continuously deactivated.

In Run 61-3 (Example III) where the hydrogen partial pressure was 270 p. s. i. g., the octane rating fell olf 0.065 number per hour showing that the process was not continuous and would require intermittent regeneration of catalyst to maintain its activity.

The data also showed that it is not possible to operate eifectively at hydrogen partial pressures above 460 for the yield in desired products decreased. Thus, the data show that hydrogen partial pressures of 450-460 p. s. i. g. are critical in that the activity of the catalyst is maintained and the yields are high. In other words, hydrogen partial pressures within the restricted range indicated, result in a process in which the yieldoctane number relationship is good, and this result is attainable in continuous on-stream operation.

The conclusion to be drawn from these runs, therefore, is clearly that there is a critical hy.. drogen partial pressure which will enable operation of the process continuously while maintaining the process so as to produce a good yield of a high octane product.

It will be understood that the foregoing examples are merely illustrative of the present invention and do not impose any limitation thereon. For example, the temperatures set forth in the above examples referred to catalyst bed temperature. This temperature range may vary within the limitations of from about 875-950 F.; a total pressure in the reaction zone may vary within the limitation of from about 600-1000. The amount of hydrogen per barrel of oil feed to the reaction zone varies according to the total pressure, and is so adjusted that the hydrogen partial pressure maintained in the reactor is of the order of about L15G-1160 p. s. i. g. It has been noted that this range is critical to secure best results. Continuous operation without catalyst regeneration has been shown to be possible at hydrogen partial pressure levels within the approximate range of from about 450-460 p. s. i. g. However, at hydrogen partial pressures above this critical pressure range of 450-460 p. s. i. g., the octane value of the product is lowered and this value decreases as the hydrogen partial pressure increases. The hydrogen production likewise decreases with increasing hydrogen partial pressures above said critical range, thus making it necessary to supply additional extraneous hydrogen to the processes, and thereby increasing the cost of producing the product. At hydrogen partial pressures below the said critical range the catalyst tends to become deactivated and this condition worsens as the hydrogen partial pressure decreases and the production phase of the process must be interrupted periodically to regenerate the catalyst. Thus, the operation should be carried out at the lowest hydrogen partial pressure commensurate with obtaining continuous operation, this hydrogen partial pressure being substantially within the said critical 450-460 p. s. i. g. range, herein disclosed.

With respect to the catalyst, although molybdenum oxide on alumina gel is preferred, many difficulty reducible oxides, such as the oxides of the 1V, VI and VIII groups of the Periodic System may be used in lieu of molybdenum oxide. Instead of using a gel alumina carrier, any active and/or high surface form of alumina may be used, such as calcined and low soda alumina, various clays, active carbon, etc., which carriers also possess good mechanical strength. The car.. rier may valso contain up to weight percent of silica, preferably in the form of a gel, where alumina constitutes the remainder of the carrier. The inclusion of 0.1 to 5% by weight of the total catalyst of boron oxide may be included in the catalyst composition for the purpose of imparting to the catalyst composition, the faculty of isomerizing the oleiins and paraflins in the original feed.

It is to be noted that although the continuous hydroforming operation given as Example I was carried out at 6500-7000 SCF/bbl. recycle rate and -75 Vol. percent hydrogen concentration in the recycle gas, these variables are not limited to these precise ranges. Recycle rates of 5500-7500 SCF/bbl. and hydrogen concentrations from 55-75 vol. percent may be used provided the operation is carried out at an appropriate pressure level between 750 and 1000 p. s. i. g. which will result in a hydrogen partial pressure in the reaction zone of I15G-4:60 p. s. i. g. It is also possible to operate at total pressures as low as 600 p. s. i. g., provided steps are taken to increase the hydrogen concentration in the recycle gas to -90% and thus produce the critical 45o-460 p. s. i. g. hydrogen partial pressure.

Numerous modications of the invention may be made without departing from the spirit thereof.

What is claimed is:

1. The method of hydroforming naphtha containing a substantial amount of naphthenic hydrocarbons which comprises feeding the said naphtha to a reaction zone simultaneously feeding to said reaction zone a hydrogen-containing gas, the amount of hydrogen-containing gas being from about 5500-7500 cu. ft. per barrel of oil, the said hydrogen-containing gas having a concentration of from i5-75% hydrogen, maintaining in said reaction zone a fluidized bed of hydroforming catalyst comprising essentially a difficulty reducible heavy metal oxide of a metal of the IV and VI groups of the Periodic System, maintaining a temperature in the said reaction zone of from about 875-950 F., maintaining a hydrogen partial pressure within the said reaction zone of from about 450-460 p. s. i. g., permitting the reactants to remain resident in the reaction'zone for a suicient period of time to effect the desired conversion and recovering from said reaction zone a product of increased aromaticity.

2. The method set forth in claim 1 in which the catalyst is molybdenum oxide on active alumina.

3. The method set forth in claim 1, characterized in that the hydroformed product of high octane value is obtained with a minimum amount of hydrogen consumption during the process by correlating the hydrogen partial pressure responsive to the amount of hydrogen fed to the reaction zone per barrel of liquid oil feed.

4. The method set forth in claim 1 in which the total pressure in the hydroforming zone is from 600 to 1000 p. s. i. g.

5. The method set forth in claim 1 in which the hydrogen-containing gas fed to the reaction zone contains up to 85-90% hydrogen.

ROBERT E. MAcPI-IERSON, Ja. OTTO SCHRICKER, JR.

References Cited in the file of this patent UNITED STATES PATENTS Number 

1. THE METHOD OF HYDROFORMING NAPHTHA CONTAINING A SUBSTANTIAL AMOUNT OF NAPHTHENIC HYDROCARBONS WHICH COMPRISES FEEDING THE SAID NAPHTHA TO REACTION ZONE SIMULTANEOUSLY FEEDING TO SAID REACTION ZONE A HYDROGEN-CONTAINING GAS, THE AMOUNT OF HYDROGEN-CONTAINING GAS BEING FROM ABOUT 5500-7500 CU. FT. PER BARREL OF OIL, THE SAID HYDROGEN-CONTAINING GAS HAVING A CONCENTRATION OF FROM 55-75% HYDROGEN, MAINTAINING IN SAID REACTION ZONE A FLUIDIZED BED OF HYDROFORMING CATALYST COMPRISING ESSENTIALLY A 